Method for cracking a hydrocarbon feedstock in a steam cracker unit

ABSTRACT

The present invention relates to a process for cracking a hydrocarbon feedstock in a steam cracker unit, comprising the following steps of:
         feeding a liquid hydrocarbon feedstock to a hydrocracking unit,   separating the stream thus hydrocracked in said hydrocracking unit into a high content aromatics stream and a gaseous stream comprising C2-C4 paraffins, hydrogen and methane,   separating C2-C4 paraffins from said gaseous stream,   feeding said C2-C4 paraffins thus separated to the furnace section of a steam cracker unit.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a continuation of U.S. application Ser. No.14/901,921, filed Dec. 29, 2015, which is a national phase under 35U.S.C. § 371 of International Application No. PCT/EP2014/063848, filedJun. 30, 2014, which claims the benefit of priority to European PatentApplication No. 13174781.8, filed Jul. 2, 2013, the entire contents ofeach of which are hereby incorporated by reference in their entirety.

BACKGROUND OF THE INVENTION

The present invention relates to a process for cracking a hydrocarbonfeedstock in a steam cracker unit.

Conventionally, crude oil is processed, via distillation, into a numberof cuts such as naphtha, gas oils and residua. Each of these cuts has anumber of potential uses such as for producing transportation fuels suchas gasoline, diesel and kerosene or as feeds to some petrochemicals andother processing units.

Light crude oil cuts such a naphthas and some gas oils can be used forproducing light olefins and single ring aromatic compounds via processessuch as steam cracking in which the hydrocarbon feed stream isevaporated and diluted with steam then exposed to a very hightemperature (800° C. to 860° C.) in short residence time (<1 second)furnace (reactor) tubes. In such a process the hydrocarbon molecules inthe feed are transformed into (on average) shorter molecules andmolecules with lower hydrogen to carbon ratios (such as olefins) whencompared to the feed molecules. This process also generates hydrogen asa useful by-product and significant quantities of lower valueco-products such as methane and C9+ Aromatics and condensed aromaticspecies (containing two or more aromatic rings which share edges).

Typically, the heavier (or higher boiling point), higher aromaticcontent streams, such as residua are further processed in a crude oilrefinery to maximize the yields of lighter (distillable) products fromthe crude oil. This processing can be carried out by processes such ashydro-cracking (whereby the hydro-cracker feed is exposed to a suitablecatalyst under conditions which result in some fraction of the feedmolecules being broken into shorter hydrocarbon molecules with thesimultaneous addition of hydrogen). Heavy refinery stream hydrocrackingis typically carried out at high pressures and temperatures and thus hasa high capital cost.

An aspect of such a combination of crude oil distillation and steamcracking of the lighter distillation cuts is the capital and other costsassociated with the fractional distillation of crude oil. Heavier crudeoil cuts (i.e. those boiling beyond ˜350° C.) are relatively rich insubstituted aromatic species and especially substituted condensedaromatic species (containing two or more aromatic rings which shareedges) and under steam cracking conditions these materials would yieldsubstantial quantities of heavy by products such as C9+ aromatics andcondensed aromatics. Hence, a consequence of the conventionalcombination of crude oil distillation and steam cracking is that asubstantial fraction of the crude oil is not processed via the steamcracker as the cracking yield of valuable products from heavier cuts isnot considered to be sufficiently high.

Another aspect of the technology discussed above is that even when onlylight crude oil cuts (such as naphtha) are processed via steam crackinga significant fraction of the feed stream is converted into low valueheavy by-products such as C9+ aromatics and condensed aromatics. Withtypical naphthas and gas oils these heavy by-products might constitute2% to 25% of the total product yield (Table VI, Page 295, Pyrolysis:Theory and Industrial Practice by Lyle F. Albright et al, AcademicPress, 1983). Whilst this represents a significant financial downgradeof expensive naphtha in lower value material on the scale of aconventional steam the yield of these heavy by-products does nottypically justify the capital investment required to up-grade thesematerials (e.g. by hydrocracking) into streams that might producesignificant quantities of higher value chemicals. This is partly becausehydrocracking plants have high capital costs and, as with mostpetrochemicals processes, the capital cost of these units typicallyscales with throughput raised to the power of 0.6 or 0.7. Consequently,the capital costs of a small scale hydro-cracking unit are normallyconsidered to be too high to justify such an investment to process steamcracker heavy by-products.

Another aspect of the conventional hydrocracking of heavy refinerystreams such as residua is that this is typically carried out undercompromise conditions chosen to achieve the desired overall conversion.As the feed streams contain a mixture of species with a range ofeasiness of cracking this result in some fraction of the distillableproducts formed by hydrocracking of relatively easily hydrocrackedspecies being further converted under the conditioned necessary tohydrocrack species more difficult to hydrocrack. This increases thehydrogen consumption and heat management difficulties associated withthe process and also increases the yield of light molecules such asmethane at the expense of more valuable species.

A feature of such a combination of crude oil distillation and steamcracking of the lighter distillation cuts is that steam cracking furnacetubes are typically unsuitable for the processing of cuts which containsignificant quantities of material with a boiling point greater than˜350° C. as it is difficult to ensure complete evaporation of these cutsprior to exposing the mixed hydrocarbon and steam stream to the hightemperatures required to promote thermal cracking. If droplets of liquidhydrocarbon are present in the hot sections of cracking tubes coke israpidly deposited on the tube surface which reduces heat transfer andincreases pressure drop and ultimately curtails the operation of thecracking tube necessitating a shut-down of the tube to allow fordecoking. Due to this difficulty a significant proportion of theoriginal crude oil cannot be processed into light olefins and aromaticspecies via a steam cracker.

US2009173665 relates to a catalyst and process for increasing themonoaromatics content of hydrocarbon feedstocks that include polynucleararomatics, wherein the increase in monoaromatics can be achieved with anincrease in gasoline/diesel yields and while reducing unwanted compoundsthereby providing a route for upgrading hydrocarbons that includesignificant quantities of polynuclear aromatics.

FR 2 364 879 relates to a selective process for producing light olefinichydrocarbons having 2 and 3 carbon atoms respectively per molecule,particularly ethylene and propylene, which are obtained byhydrogenolysis or hydrocracking followed with steam-cracking.

DE 2708412 relates to an integrated process for producing ethylenecomprising a step of introducing a hydrocarbon feed stock in ahydrocracking reactor free of non-thermal hydrocracking catalyst underconditions of temperature in the range of from 510-815° C., a pressurein the range of from 15-70 atm and a residence time in the range of from5-60 sec, a step of separating aromatic hydrocarbons, a step ofseparating a stream consisting essentially of hydrocarbons C2-C3 bycryogenic techniques; and introducing said stream into a cracking zonemaintained under conditions for converting hydrocarbons in a streammainly comprising ethylene.

U.S. Pat. No. 3,944,481 relates to a process for converting crude oilfractions into an olefin product by hydrocracking the crude oilfractions to C2-C5 paraffins and thermally cracking these to a C2-C3olefin mixture, wherein hydrogen, methane, and the C6-C9 components areseparated from the C2-C5 fraction and the C2-C5 fraction is then mixedwith steam and injected into the thermal cracker.

EP 0023802 relates to a process for producing light paraffins whichcomprises hydrocracking a crude oil fraction boiling in the range 93° C.to 538° C. at a pressure above 2859 kPa, a temperature of 300 to 565°C., a hydrogen/hydrocarbon mole ratio of 4:1 to 50:1 and a residencetime of 1 to 180 seconds over a catalyst, and recovering a C2-C5 alkaneproduct, wherein said alkane product is thermally cracked to ethyleneand propylene.

GB 1148967 relates to a process for the preparation of ethylene,comprising hydrocracking a hydrocarbon oil boiling below 250° C. atelevated temperature in the presence of hydrogen and a catalyst,separating a hydrocarbon mixture from the reaction product of thehydrocracking, thermally cracking this mixture at an elevatedtemperature in the presence of steam, and separating ethylene from thereaction product of the thermal cracking. The hydrocarbon mixture formedfrom the hydrocarbon oil as a result of the hydrocracking is present asa liquid in the gas/liquid separating system and is passed through aline to the steam cracking furnace.

GB 1250615 relates to an aromatic extraction process whereby aromaticscan be extracted from aromatic-containing hydro crackates

U.S. Pat. No. 3,842,138 relates to a method of thermal cracking in thepresence of hydrogen of a charge of hydrocarbons of petroleum whereinthe hydrocracking process is carried out under a pressure of 5 and 70bars at the outlet of the reactor with very short residence times of0.01 and 0.5 second and a temperature range at the outlet of the reactorextending from 625 to 1000° C.

An object of the present invention is to provide a method for upgradingnaphtha to aromatics and steam cracker feedstock comprising C2-C4paraffins.

Another object of the present invention is to provide a method forconverting relatively heavy liquid feeds, such as diesel and atmosphericgasoil to produce a hydrocracking product stream comprisingmono-aromatic hydrocarbons and C2-C4 paraffins.

Another object of the present invention is to process a heavy liquidfeedstock while minimizing the production of heavy C9+ byproducts.

The present invention relates to a process for cracking a hydrocarbonfeedstock in a steam cracker unit, comprising the following steps of:

feeding a liquid hydrocarbon feedstock to a hydrocracking unit,

separating the stream thus hydrocracked in said hydrocracking unit intoa high content aromatics stream and a gaseous stream comprising C2-C4paraffins, hydrogen and methane,

separating C2-C4 paraffins from said gaseous stream,

feeding said C2-C4 paraffins thus separated to the furnace section of asteam cracker unit.

On basis of such a process one or more of the presents objects areachieved.

According to such a method the hydrocracked feedstock can be used as afeedstock for a steam cracker unit. The aromatics get separated from thegaseous stream and will not be sent to the steam cracker unit but willbe further processed in a separate unit. Methane and other lightcomponents can be pretreated in a separator to separate C2-C4 paraffinsfrom said gaseous stream and the C2-C4 paraffins be sent to a steamcracker unit. The typical gas produced in a steam cracker unit usuallycontains a lot of hydrogen often fueled to satisfy the energy demand ofthe steam cracker. This therefore will improve the energy performanceand hydrogen balance.

Consequently the main aim of the present invention is to produce LPG(Ethane, propane, butanes), which can be processed in a gas steamcracker unit. By placing a suitably designed hydrocracker unit in frontof a gas steam cracker, and feeding said hydrocracker unit with a liquidhydrocarbon feedstock, such as for instance naphtha, it effectively actsas a feed pretreatment unit to the gas steam cracker making an otherwiseunsuitable feed suitable for processing in existing gas steam crackers.Potential feeds to be processed according to the present method could behydrocarbon streams such as diesel, kerosene, atmospheric gasoils (AGO),gas condensates, naphtha and waxy materials.

According to the present invention the process for producing C2-C4olefins and aromatic hydrocarbons having one aromatic ring comprises astep of contacting a liquid hydrocarbon feedstock having a boiling pointin the range of 20-350° C. in the presence of hydrogen with ahydrocracking catalyst to produce a hydrocracking product streamcomprising aromatic hydrocarbons having one aromatic ring and C2-C4paraffins, a step of separating the aromatic hydrocarbons having onearomatic ring from the hydrocracking product stream; and a step ofseparating the C2-C4 paraffins from the hydrocracking product stream andpreferably feeding said separated C2-C4 paraffins into dedicatedpyrolysis furnaces to conduct a pyrolysis reaction to produce apyrolysis product stream comprising C2-C4 olefins.

In preferred embodiments a portion of the hydrogen comprised in thepyrolysis product stream is separated and fed to the hydrocracking step,C5+ hydrocarbons comprised in the pyrolysis product stream are separatedand are fed to the hydrocracking step, aromatic hydrocarbons having onearomatic ring are separated from the hydrocracking product stream usingseparator, e.g. a distillation process or a solvent extraction process,C2-C4 paraffins are separated from the hydrocracking product streamgaseous fraction using a separator, e.g. a distillation process.

According to other preferred embodiments one can pretreat thehydrocarbon feed (e.g. by solvent extraction) to separate thearomatics+naphtenes (extract) and the paraffins (raffinate).Consequently, one could process the aromatics+naphtenes stream via ahydrocracker (to do the HDS and remove any traces of paraffins anddehydrogenate the naphtenes to make sales specification aromatics) andsend the paraffins to the steam cracker to make light olefins. Thiscombination of processes will minimize the hydrogen consumption andminimize fuel gas make from the steam cracker unit.

As discussed above, the present process further comprises separatingsaid high content aromatics stream into a stream of heavy aromatics anda stream high in mono-aromatics, especially using a separator of thedistillation type.

According to another embodiment of the present invention the presentprocess further comprises feeding the gaseous stream, i.e. predominantlycomprising C2-C4 paraffins, to a dehydrogenation unit for obtaininghydrogen, C3-olefins and C4-olefins. This means that the gaseous streamproduced in the hydrocracking unit can be sent to different routes, thatis to the steam cracker unit or to the dehydrogenation unit. This choiceof routes provides flexibility to the present method. Processes for thedehydrogenation of lower alkanes such as propane and butanes aredescribed as lower alkane dehydrogenation process.

In a preferred embodiment the present process further comprisesseparating C2-C4 paraffins into individual streams, each streampredominantly comprising C2 paraffins, C3 paraffins and C4 paraffins,respectively, and feeding each individual stream to a specific furnacesection of said steam cracker unit. This means that a C2 stream, whichstream predominantly comprises C2 paraffins, is sent to a specific C2furnace section of the steam cracker unit. The same holds for a C3stream and a C4 stream. Such a separation into specific streams has apositive influence on the product yield of the steam cracker unit. In apreferred embodiment the C3 stream and the C4 stream, as separatestreams or as a combined C3+C4 stream, are sent to a dehydrogenationunit. More in detail, the present method further comprises separatingC2-C4 paraffins into individual streams, each stream predominantlycomprising C2 paraffins, C3 paraffins and C4 paraffins, respectively,and feeding the predominantly comprising C2 paraffin stream to a steamcracker whilst feeding the predominantly comprising C3 paraffin streamto a propane dehydrogenation unit and feeding the predominantlycomprising C4 paraffin stream to a butane dehydrogenation unit.

As discussed above, the gaseous stream produced in the hydrocrackingunit contains a broad spectrum of hydrocracked products. In a preferredembodiment not only the C2-C4 paraffins will be recovered from thegaseous product but the other valuable components, such as hydrogen andmethane as well. The hydrogen and methane containing stream willpreferably be recycled to the hydrocracking unit. In addition, it isalso preferred to have a purge stream in the hydrocracking unit toprevent accumulation of unwanted components.

The method for separating said C2-C4 paraffins from said gaseous streamis preferably carried out by a separation of the type chosen fromcryogenic distillation or solvent extraction.

The preferred process conditions in said hydrocracking include atemperature of 300-550° C., a pressure of 300-5000 kPa gauge and aWeight Hourly Space Velocity of 0.1-10 h−1. More preferred hydrocrackingconditions include a temperature of 350-550° C., a pressure of 600-3000kPa gauge and a Weight Hourly Space Velocity of 0.2-2 h−1.

The reactor type design of the said hydrocracking unit is chosen fromthe group of the fixed bed type, ebulating bed reactor type and theslurry type, wherein the fixed bed type is preferred.

Examples of the hydrocarbon feedstock to said hydrocracking unit are ofthe type naphtha, kerosene, diesel, atmospheric gas oil (AGO), gascondensates, naphtha, waxes, or combinations thereof.

The separation of the high content aromatics stream is preferably of thedistillation type.

According to a preferred embodiment the process according to the presentinvention further comprises separating C7 to C9 aromatics, for exampletoluene and xylene rich fraction, from the stream high inmono-aromatics, and converting said C7 to C9 aromatics into a benzenerich fraction.

Moreover, the process according to the present invention furthercomprises recovering a portion of hydrogen from the product stream ofsaid steam cracker unit and feeding said hydrogen to said hydrocrackingunit.

As discussed above, the present method further comprises recovering C5+hydrocarbons from the product stream of said steam cracker unit andfeeding said C5+ hydrocarbons to said hydrocracking unit.

The term “crude oil” as used herein refers to the petroleum extractedfrom geologic formations in its unrefined form. Any crude oil issuitable as the source material for the process of this invention,including Arabian Heavy, Arabian Light, other Gulf crudes, Brent, NorthSea crudes, North and West African crudes, Indonesian, Chinese crudesand mixtures thereof, but also shale oil, tar sands and bio-based oils.The crude oil is preferably a conventional petroleum having an APIgravity of more than 20° API as measured by the ASTM D287 standard. Morepreferably, the crude oil used is a light crude oil having an APIgravity of more than 30° API. Most preferably, the crude oil comprisesArabian Light Crude Oil. Arabian Light Crude Oil typically has an APIgravity of between 32-36° API and a sulfur content of between 1.5-4.5wt-%.

The term “petrochemicals” or “petrochemical products” as used hereinrelates to chemical products derived from crude oil that are not used asfuels. Petrochemical products include olefins and aromatics that areused as a basic feedstock for producing chemicals and polymers.High-value petrochemicals include olefins and aromatics. Typicalhigh-value olefins include, but are not limited to, ethylene, propylene,butadiene, butylene-1, isobutylene, isoprene, cyclopentadiene andstyrene. Typical high-value aromatics include, but are not limited to,benzene, toluene, xylene and ethyl benzene.

The term “fuels” as used herein relates to crude oil-derived productsused as energy carrier. Unlike petrochemicals, which are a collection ofwell-defined compounds, fuels typically are complex mixtures ofdifferent hydrocarbon compounds. Fuels commonly produced by oilrefineries include, but are not limited to, gasoline, jet fuel, dieselfuel, heavy fuel oil and petroleum coke.

The term “aromatic hydrocarbons” or “aromatics” is very well known inthe art. Accordingly, the term “aromatic hydrocarbon” relates tocyclically conjugated hydrocarbon with a stability (due todelocalization) that is significantly greater than that of ahypothetical localized structure (e.g. Kekulé structure). The mostcommon method for determining aromaticity of a given hydrocarbon is theobservation of diatropicity in the 1H NMR spectrum, for example thepresence of chemical shifts in the range of from 7.2 to 7.3 ppm forbenzene ring protons. The terms “naphthenic hydrocarbons” or“naphthenes” or “cycloalkanes” is used herein having its establishedmeaning and accordingly relates types of alkanes that have one or morerings of carbon atoms in the chemical structure of their molecules.

The term “olefin” is used herein having its well-established meaning.Accordingly, olefin relates to an unsaturated hydrocarbon compoundcontaining at least one carbon-carbon double bond. Preferably, the term“olefins” relates to a mixture comprising two or more of ethylene,propylene, butadiene, butylene-1, isobutylene, isoprene andcyclopentadiene.

The term “LPG” as used herein refers to the well-established acronym forthe term “liquefied petroleum gas”. LPG generally consists of a blend ofC2-C4 hydrocarbons i.e. a mixture of C2, C3, and C4 hydrocarbons.

The term “BTX” as used herein relates to a mixture of benzene, tolueneand xylenes.

As used herein, the term “C # hydrocarbons”, wherein “#” is a positiveinteger, is meant to describe all hydrocarbons having # carbon atoms.Moreover, the term “C #+ hydrocarbons” is meant to describe allhydrocarbon molecules having # or more carbon atoms. Accordingly, theterm “C5+ hydrocarbons” is meant to describe a mixture of hydrocarbonshaving 5 or more carbon atoms. The term “C5+ alkanes” accordinglyrelates to alkanes having 5 or more carbon atoms.

As used herein, the term “hydrocracker unit” or “hydrocracker” relatesto a refinery unit in which a hydrocracking process is performed i.e. acatalytic cracking process assisted by the presence of an elevatedpartial pressure of hydrogen; see e.g. Alfke et al. (2007) loc.cit. Theproducts of this process are saturated hydrocarbons and, depending onthe reaction conditions such as temperature, pressure and space velocityand catalyst activity, aromatic hydrocarbons including BTX. The processconditions used for hydrocracking generally includes a processtemperature of 200-600° C., elevated pressures of 0.2-20 MPa, spacevelocities between 0.1-10 h−1

Hydrocracking reactions proceed through a bifunctional mechanism whichrequires a acid function, which provides for the cracking andisomerization and which provides breaking and/or rearrangement of thecarbon-carbon bonds comprised in the hydrocarbon compounds comprised inthe feed, and a hydrogenation function. Many catalysts used for thehydrocracking process are formed by composting various transitionmetals, or metal sulfides with the solid support such as alumina,silica, alumina-silica, magnesia and zeolites.

As used herein, the term “feed hydrocracking unit” or “FHC” refers to arefinery unit for performing a hydrocracking process suitable forconverting a complex hydrocarbon feed that is relatively rich innaphthenic and paraffinic hydrocarbon compounds—such as straight runcuts including, but not limited to, naphtha- to LPG and alkanes.Preferably, the hydrocarbon feed that is subject to feed hydrocrackingcomprises naphtha. Accordingly, the main product produced by feedhydrocracking is LPG that is to be converted into olefins (i.e. to beused as a feed for the conversion of alkanes to olefins). The FHCprocess may be optimized to keep one aromatic ring intact of thearomatics comprised in the FHC feedstream, but to remove most of theside-chains from said aromatic ring. In such a case, the processconditions to be employed for FHC are comparable to the processconditions to be used in the GHC process as described herein above.Alternatively, the FHC process can be optimized to open the aromaticring of the aromatic hydrocarbons comprised in the FHC feedstream. Thiscan be achieved by modifying the GHC process as described herein byincreasing the hydrogenation activity of the catalyst, optionally incombination with selecting a lower process temperature, optionally incombination with a reduced space velocity. In such a case, preferredfeed hydrocracking conditions thus include a temperature of 300-550° C.,a pressure of 300-5000 kPa gauge and a Weight Hourly Space Velocity of0.1-10 h−1. More preferred feed hydrocracking conditions include atemperature of 300-450° C., a pressure of 300-5000 kPa gauge and aWeight Hourly Space Velocity of 0.1-10 h−1. Even more preferred FHCconditions optimized to the ring-opening of aromatic hydrocarbonsinclude a temperature of 300-400° C., preferably a temperature of350-450° C., more preferably 375-450° C., a pressure of 600-3000 kPagauge and a Weight Hourly Space Velocity of 0.2-2 h−1.

The “aromatic ring opening unit” refers to a refinery unit wherein thearomatic ring opening process is performed. Aromatic ring opening is aspecific hydrocracking process that is particularly suitable forconverting a feed that is relatively rich in aromatic hydrocarbon havinga boiling point in the kerosene and gasoil boiling point range toproduce LPG and, depending on the process conditions, a light-distillate(ARO-derived gasoline). Such an aromatic ring opening process (AROprocess) is for instance described in U.S. Pat. Nos. 3,256,176 and4,789,457. Such processes may comprise of either a single fixed bedcatalytic reactor or two such reactors in series together with one ormore fractionation units to separate desired products from unconvertedmaterial and may also incorporate the ability to recycle unconvertedmaterial to one or both of the reactors. Reactors may be operated at atemperature of 200-600° C., preferably 300-400° C., a pressure of 3-35MPa, preferably 5 to 20 MPa together with 5-20 wt-% of hydrogen (inrelation to the hydrocarbon feedstock), wherein said hydrogen may flowco-current with the hydrocarbon feedstock or counter current to thedirection of flow of the hydrocarbon feedstock, in the presence of adual functional catalyst active for both hydrogenation-dehydrogenationand ring cleavage, wherein said aromatic ring saturation and ringcleavage may be performed. Catalysts used in such processes comprise oneor more elements selected from the group consisting of Pd, Rh, Ru, Ir,Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V in metallic or metalsulphide form supported on an acidic solid such as alumina, silica,alumina-silica and zeolites. In this respect, it is to be noted that theterm “supported on” as used herein includes any conventional way toprovide a catalyst which combines one or more elements with a catalyticsupport. A further aromatic ring opening process (ARO process) isdescribed in U.S. Pat. No. 7,513,988. Accordingly, the ARO process maycomprise aromatic ring saturation at a temperature of 100-500° C.,preferably 200-500° C. and more preferably 300-500° C., a pressure of2-10 MPa together with 5-30 wt-%, preferably 10-30 wt-% of hydrogen (inrelation to the hydrocarbon feedstock) in the presence of an aromatichydrogenation catalyst and ring cleavage at a temperature of 200-600°C., preferably 300-400° C., a pressure of 1-12 MPa together with 5-20wt-% of hydrogen (in relation to the hydrocarbon feedstock) in thepresence of a ring cleavage catalyst, wherein said aromatic ringsaturation and ring cleavage may be performed in one reactor or in twoconsecutive reactors. The aromatic hydrogenation catalyst may be aconventional hydrogenation/hydrotreating catalyst such as a catalystcomprising a mixture of Ni, W and Mo on a refractory support, typicallyalumina. The ring cleavage catalyst comprises a transition metal ormetal sulphide component and a support. Preferably the catalystcomprises one or more elements selected from the group consisting of Pd,Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W and V in metallicor metal sulphide form supported on an acidic solid such as alumina,silica, alumina-silica and zeolites. By adapting either single or incombination the catalyst composition, operating temperature, operatingspace velocity and/or hydrogen partial pressure, the process can besteered towards full saturation and subsequent cleavage of all rings ortowards keeping one aromatic ring unsaturated and subsequent cleavage ofall but one ring. In the latter case, the ARO process produces alight-distillate (“ARO-gasoline”) which is relatively rich inhydrocarbon compounds having one aromatic ring.

As used herein, the term “dearomatization unit” relates to a refineryunit for the separation of aromatic hydrocarbons, such as BTX, from amixed hydrocarbon feed. Such dearomatization processes are described inFolkins (2000) Benzene, Ullmann's Encyclopedia of Industrial Chemistry.Accordingly, processes exist to separate a mixed hydrocarbon stream intoa first stream that is enriched for aromatics and a second stream thatis enriched for paraffins and naphthenes. A preferred method to separatearomatic hydrocarbons from a mixture of aromatic and aliphatichydrocarbons is solvent extraction; see e.g. WO 2012135111 A2. Thepreferred solvents used in aromatic solvent extraction are sulfolane,tetraethylene glycol and N-methylpyrolidone which are commonly usedsolvents in commercial aromatics extraction processes. These species areoften used in combination with other solvents or other chemicals(sometimes called co-solvents) such as water and/or alcohols.Non-nitrogen containing solvents such as sulfolane are particularlypreferred. Commercially applied dearomatization processes are lesspreferred for the dearomatization of hydrocarbon mixtures having aboiling point range that exceeds 250° C., preferably 200° C., as theboiling point of the solvent used in such solvent extraction needs to belower than the boiling point of the aromatic compounds to be extracted.Solvent extraction of heavy aromatics is described in the art; see e.g.U.S. Pat. No. 5,880,325. Alternatively, other known methods than solventextraction, such as molecular sieve separation or separation based onboiling point, can be applied for the separation of heavy aromatics in adearomatization process.

A process to separate a mixed hydrocarbon stream into a streamcomprising predominantly paraffins and a second stream comprisingpredominantly aromatics and naphthenes comprises processing said mixedhydrocarbon stream in a solvent extraction unit comprising three mainhydrocarbon processing columns: solvent extraction column, strippercolumn and extract column. Conventional solvents selective for theextraction of aromatics are also selective for dissolving lightnaphthenic and to a lesser extent light paraffinic species hence thestream exiting the base of the solvent extraction column comprisessolvent together with dissolved aromatic, naphthenic and lightparaffinic species. The stream exiting the top of the solvent extractioncolumn (often termed the raffinate stream) comprises the relativelyinsoluble, with respect to the chosen solvent) paraffinic species. Thestream exiting the base of the solvent extraction column is thensubjected, in a distillation column, to evaporative stripping in whichspecies are separated on the basis of their relative volatility in thepresence of the solvent. In the presence of a solvent, light paraffinicspecies have higher relative volatilities than naphthenic species andespecially aromatic species with the same number of carbon atoms, hencethe majority of light paraffinic species may be concentrated in theoverhead stream from the evaporative stripping column. This stream maybe combined with the raffinate stream from the solvent extraction columnor collected as a separate light hydrocarbon stream. Due to theirrelatively low volatility the majority of the naphthenic and especiallyaromatic species are retained in the combined solvent and dissolvedhydrocarbon stream exiting the base of this column. In the finalhydrocarbon processing column of the extraction unit, the solvent isseparated from the dissolved hydrocarbon species by distillation. Inthis step the solvent, which has a relatively high boiling point, isrecovered as the base stream from the column whilst the dissolvedhydrocarbons, comprising mainly aromatics and naphthenic species, arerecovered as the vapour stream exiting the top of the column. Thislatter stream is often termed the extract.

The process of the present invention may require removal of sulfur fromcertain crude oil fractions to prevent catalyst deactivation indownstream refinery processes, such as catalytic reforming or fluidcatalytic cracking. Such a hydrodesulfurization process is performed ina “HDS unit” or “hydrotreater”; see Alfke (2007) loc. cit. Generally,the hydrodesulfurization reaction takes place in a fixed-bed reactor atelevated temperatures of 200-425° C., preferably of 300-400° C. andelevated pressures of 1-20 MPa gauge, preferably 1-13 MPa gauge in thepresence of a catalyst comprising elements selected from the groupconsisting of Ni, Mo, Co, W and Pt, with or without promoters, supportedon alumina, wherein the catalyst is in a sulfide form.

In a further embodiment, the process further comprises ahydrodealkylation step wherein the BTX (or only the toluene and xylenesfraction of said BTX produced) is contacted with hydrogen underconditions suitable to produce a hydrodealkylation product streamcomprising benzene and fuel gas.

The process step for producing benzene from BTX may include a stepwherein the benzene comprised in the hydrocracking product stream isseparated from the toluene and xylenes before hydrodealkylation. Theadvantage of this separation step is that the capacity of thehydrodealkylation reactor is increased. The benzene can be separatedfrom the BTX stream by conventional distillation.

Processes for hydrodealkylation of hydrocarbon mixtures comprising C6-C9aromatic hydrocarbons are well known in the art and include thermalhydrodealkylation and catalytic hydrodealkylation; see e.g. WO2010/102712 A2. Catalytic hydrodealkylation is preferred as thishydrodealkylation process generally has a higher selectivity towardsbenzene than thermal hydrodealkylation. Preferably catalytichydrodealkylation is employed, wherein the hydrodealkylation catalyst isselected from the group consisting of supported chromium oxide catalyst,supported molybdenum oxide catalyst, platinum on silica or alumina andplatinum oxide on silica or alumina.

The process conditions useful for hydrodealkylation, also describedherein as “hydrodealkylation conditions”, can be easily determined bythe person skilled in the art. The process conditions used for thermalhydrodealkylation are for instance described in DE 1668719 A1 andinclude a temperature of 600-800° C., a pressure of 3-10 MPa gauge and areaction time of 15-45 seconds. The process conditions used for thepreferred catalytic hydrodealkylation are described in WO 2010/102712 A2and preferably include a temperature of 500-650° C., a pressure of 3.5-8MPa gauge, preferably of 3.5-7 MPa gauge and a Weight Hourly SpaceVelocity of 0.5-2 h−1. The hydrodealkylation product stream is typicallyseparated into a liquid stream (containing benzene and other aromaticsspecies) and a gas stream (containing hydrogen, H2S, methane and otherlow boiling point hydrocarbons) by a combination of cooling anddistillation. The liquid stream may be further separated, bydistillation, into a benzene stream, a C7 to C9 aromatics stream andoptionally a middle-distillate stream that is relatively rich inaromatics. The C7 to C9 aromatic stream may be fed back to reactorsection as a recycle to increase overall conversion and benzene yield.The aromatic stream which contains polyaromatic species such asbiphenyl, is preferably not recycled to the reactor but may be exportedas a separate product stream and recycled to the integrated process asmiddle-distillate (“middle-distillate produced by hydrodealkylation”).The gas stream contains significant quantities of hydrogen may berecycled back the hydrodealkylation unit via a recycle gas compressor orto any other refinery that uses hydrogen as a feed. A recycle gas purgemay be used to control the concentrations of methane and H2S in thereactor feed.

As used herein, the term “gas separation unit” relates to the refineryunit that separates different compounds comprised in the gases producedby the crude distillation unit and/or refinery unit-derived gases.Compounds that may be separated to separate streams in the gasseparation unit comprise ethane, propane, butanes, hydrogen and fuel gasmainly comprising methane. Any conventional method suitable for theseparation of said gases may be employed. Accordingly, the gases may besubjected to multiple compression stages wherein acid gases such as CO2and H2S may be removed between compression stages. In a following step,the gases produced may be partially condensed over stages of a cascaderefrigeration system to about where only the hydrogen remains in thegaseous phase. The different hydrocarbon compounds may subsequently beseparated by distillation.

A process for the conversion of alkanes to olefins involves “steamcracking” or “pyrolysis”. As used herein, the term “steam cracking”relates to a petrochemical process in which saturated hydrocarbons arebroken down into smaller, often unsaturated, hydrocarbons such asethylene and propylene. In steam cracking gaseous hydrocarbon feeds likeethane, propane and butanes, or mixtures thereof, (gas cracking) orliquid hydrocarbon feeds like naphtha or gasoil (liquid cracking) isdiluted with steam and briefly heated in a furnace without the presenceof oxygen. Typically, the reaction temperature is 750-900° C., but thereaction is only allowed to take place very briefly, usually withresidence times of 50-1000 milliseconds. Preferably, a relatively lowprocess pressure is to be selected of atmospheric up to 175 kPa gauge.Preferably, the hydrocarbon compounds ethane, propane and butanes areseparately cracked in accordingly specialized furnaces to ensurecracking at optimal conditions. After the cracking temperature has beenreached, the gas is quickly quenched to stop the reaction in a transferline heat exchanger or inside a quenching header using quench oil. Steamcracking results in the slow deposition of coke, a form of carbon, onthe reactor walls. Decoking requires the furnace to be isolated from theprocess and then a flow of steam or a steam/air mixture is passedthrough the furnace coils. This converts the hard solid carbon layer tocarbon monoxide and carbon dioxide. Once this reaction is complete, thefurnace is returned to service. The products produced by steam crackingdepend on the composition of the feed, the hydrocarbon to steam ratioand on the cracking temperature and furnace residence time. Lighthydrocarbon feeds such as ethane, propane, butane or light naphtha giveproduct streams rich in the lighter polymer grade olefins, includingethylene, propylene, and butadiene. Heavier hydrocarbon (full range andheavy naphtha and gas oil fractions) also give products rich in aromatichydrocarbons.

To separate the different hydrocarbon compounds produced by steamcracking the cracked gas is subjected to a fractionation unit. Suchfractionation units are well known in the art and may comprise aso-called gasoline fractionator where the heavy-distillate (“carbonblack oil”) and the middle-distillate (“cracked distillate”) areseparated from the light-distillate and the gases. In the subsequentoptional quench tower, most of the light-distillate produced by steamcracking (“pyrolysis gasoline” or “pygas”) may be separated from thegases by condensing the light-distillate. Subsequently, the gases may besubjected to multiple compression stages wherein the remainder of thelight distillate may be separated from the gases between the compressionstages. Also acid gases (CO2 and H2S) may be removed between compressionstages. In a following step, the gases produced by pyrolysis may bepartially condensed over stages of a cascade refrigeration system toabout where only the hydrogen remains in the gaseous phase. Thedifferent hydrocarbon compounds may subsequently be separated by simpledistillation, wherein the ethylene, propylene and C4 olefins are themost important high-value chemicals produced by steam cracking. Themethane produced by steam cracking is generally used as fuel gas, thehydrogen may be separated and recycled to processes that consumehydrogen, such as hydrocracking processes. The acetylene produced bysteam cracking preferably is selectively hydrogenated to ethylene. Thealkanes comprised in the cracked gas may be recycled to the process forolefins synthesis.

The term “propane dehydrogenation unit” as used herein relates to apetrochemical process unit wherein a propane feedstream is convertedinto a product comprising propylene and hydrogen. Accordingly, the term“butane dehydrogenation unit” relates to a process unit for converting abutane feedstream into C4 olefins. Together, processes for thedehydrogenation of lower alkanes such as propane and butanes aredescribed as lower alkane dehydrogenation process. Processes for thedehydrogenation of lower alkanes are well-known in the art and includeoxidative dehydrogenation processes and non-oxidative dehydrogenationprocesses. In an oxidative dehydrogenation process, the process heat isprovided by partial oxidation of the lower alkane(s) in the feed. In anon-oxidative dehydrogenation process, which is preferred in the contextof the present invention, the process heat for the endothermicdehydrogenation reaction is provided by external heat sources such ashot flue gases obtained by burning of fuel gas or steam. In anon-oxidative dehydrogenation process the process conditions generallycomprise a temperature of 540-700° C. and an absolute pressure of 25-500kPa. For instance, the UOP Oleflex process allows for thedehydrogenation of propane to form propylene and of (iso)butane to form(iso)butylene (or mixtures thereof) in the presence of a catalystcontaining platinum supported on alumina in a moving bed reactor; seee.g. U.S. Pat. No. 4,827,072. The Uhde STAR process allows for thedehydrogenation of propane to form propylene or of butane to formbutylene in the presence of a promoted platinum catalyst supported on azinc-alumina spinel; see e.g. U.S. Pat. No. 4,926,005. The STAR processhas been recently improved by applying the principle ofoxydehydrogenation. In a secondary adiabatic zone in the reactor part ofthe hydrogen from the intermediate product is selectively converted withadded oxygen to form water. This shifts the thermodynamic equilibrium tohigher conversion and achieves a higher yield. Also the external heatrequired for the endothermic dehydrogenation reaction is partly suppliedby the exothermic hydrogen conversion. The Lummus Catofin processemploys a number of fixed bed reactors operating on a cyclical basis.The catalyst is activated alumina impregnated with 18-20 wt-% chromium;see e.g. EP 0 192 059 A1 and GB 2 162 082 A. The Catofin process has theadvantage that it is robust and capable of handling impurities whichwould poison a platinum catalyst. The products produced by a butanedehydrogenation process depends on the nature of the butane feed and thebutane dehydrogenation process used. Also the Catofin process allows forthe dehydrogenation of butane to form butylene; see e.g. U.S. Pat. No.7,622,623.

The present invention will be discussed in the next Example whichexample should not be interpreted as limiting the scope of protection.

The sole FIGURE provides a schematic flow sheet of an embodiment of thepresent invention.

EXAMPLES

Feedstock 33, which can include different types of feedstock, forexample naphtha 35, kerosene 36, diesel 37, atmospheric gas oil (AGO) 38originating from tanks 2, 3, 4, 5 respectively, is sent to ahydrocracker unit 17. In hydrocracking unit 17 a feedstock 33 ishydrocracked in the presence of hydrogen. The hydrocracking processresults in the formation of a gaseous stream 19 comprising C2-C4paraffins, hydrogen and methane and a high content aromatics stream 40.The gaseous stream 19 is sent to a separator 12, e.g. cryogenicdistillation or solvent extraction, and separated into differentstreams, i.e. a stream 55 comprising C2-C4 paraffins, a stream 20comprising hydrogen and methane and a purge stream 18. Stream 20 can berecycled to hydrocracking unit 17.

As mentioned before, stream 55 can be sent directly (not shown) toeither a dehydrogenation unit 57 or directly (not shown) to a steamcracker unit 11. However, before sending stream 55 to steam cracker unit11 it is preferred to carry out a separation on stream 55 first. Inseparator 56 the C2-C4 paraffins are separated into individual streams30, 31 and 32. This means that stream 30 predominantly comprises C2paraffins, stream 31 predominantly comprises C3 paraffins and stream 32predominantly comprises C4 paraffins. If necessary, further separationof unwanted components or temperature adjustments can made in units 9,10, 13. The individual streams 21, 27 and 29 will be sent to specificfurnace sections of steam cracker unit 11. Although steam cracker unit11 is shown as one single unit, in the present method it is to beunderstood that in a preferred embodiment steam cracker unit 11comprises different furnace sections each dedicated for a specificchemical composition, that is a furnace section for C2, a furnacesection for C3 and a furnace section for C4. In a preferred embodimentstream 27 predominantly comprising C3 paraffins and stream 29predominantly comprising C4 paraffins are sent as stream 54 and stream23 to dehydrogenation unit 57, respectively. In another embodiment it ispossible to separate only C3 and C4 from stream 55 and to send acombined C3 and C4 stream to dehydrogenation unit 57.

In steam cracker unit 11 streams 21, 27 and 29 and a feedstock 58, forexample C2 to C4 gases coming from an unit 1, are processed and itsreaction products 28 are separated in a separation section 6. A C2-C6alkanes comprising gas stream 7 is recycled to the steam cracker unit11. Hydrogen 15 and pygas 14 can be sent to hydrocracking unit 17. Thevaluable product stream 8 comprising unsaturated hydrocarbons such aslighter alkenes including ethylene, propylene and butadienes is sent tofurther petrochemical processes. In case heavy hydrocarbons such ascarbon black oil, cracked distillate and C9+ hydrocarbons are producedin steam cracker unit 11 these products can optionally be recycled tohydrocracking unit 17 as well.

High content aromatics stream 40 is sent to a separator 16, for examplea distillation process, and separated into a stream 41 of heavyaromatics and a stream 43 high in mono-aromatics. Stream 42 whichpredominantly comprises C7 to C9 aromatics can converted in unit 24 intoa benzene rich fraction 59 and a methane rich fraction 44.

The Example disclosed herein makes a distinction between a process(case 1) in which the naphtha is only processed through a steam crackerunit and a process (case 2) wherein the naphtha is sent to ahydrocracking unit, wherein in the gaseous stream thus formed the C2-C4paraffins is separated and fed to the furnace section of a steam crackerunit. Case 1 is a comparative example and case 2 is an example accordingto the present invention.

The conditions for the steam cracker are as follows: Ethane and Propanefurnaces: Coil Outlet temperature=845° C., Steam-to-oil-ratio=0.37,C4-furnaces: Coil Outlet temperature=820° C., Steam-to-oil-ratio=0.37,Liquid furnaces: Coil Outlet temperature=820° C.,Steam-to-oil-ratio=0.37. Regarding the specific conditions for theHydrocracking unit 17: The modeling was carried out for a hydrocrackingreactor operating conditions: mean reactor temperature 510° C., Weighthour space velocity of 1 hr-1 and a reactor pressure of 1379 kPa gauge.The catalyst comprised a mixture of Pt supported on gamma alumina andHZSM-5 with a Si:Al ratio of 100:1.

The composition of the naphtha feed can be found in Table 1.

TABLE 1 Composition of naphtha as feedstock Naphtha n-Paraffins wt-%36.7 i-Paraffins wt-% 38.2 Naphthenes wt-% 20.1 Aromatics wt-% 5.0Density 60 F. kg/L 0.673 IBP ° C. 37.8 BP10 ° C. 45.7 BP30 ° C. 49.5BP50 ° C. 54.7 BP70 ° C. 64.0 BP90 ° C. 79.4 FBP ° C. 103.0

The battery product slate (wt. % of feed) for each of the case 1 andcase 2 can be found in Table 2.

TABLE 2 battery product slate (wt. % of feed) Feed: naphtha CASE 1 CASE2 BATTERY LIMIT PRODUCT SLATE SC FHC + SC H2 0.9 0.5 CH4 17.9 18.3ETHYLENE 35.1 50.1 PROPYLENE 19.3 12.8 BUTADIENE 5.3 2.0 ISO-BUTENE 3.30.3 BENZENE 9.1 8.3 TX CUT 3.4 7.0 STYRENE 0.9 0.1 OTHER C7-C8 0.4 0.0C9 RESIN FEED 0.7 0.0 CD 1.5 0.2 CBO 2.0 0.3 % HIGH VALUE CHEMICALS 76.980.5

From Table 2 one can see that treating the naphtha in a hydrocrackingunit according to the present method (case 2) causes an increase oftotal BTX (benzene, toluene plus xylenes. Accordingly, the resultsdisclosed in Table 2 show a considerable increase in BTX from case 1(comparative example) to case 2 (according to the present invention). CDmeans cracked distillate and CBO means carbon black oil, respectively.Table 2 further illustrates that the yield of high value chemicals(ethylene+propylene+butadiene+iso-butene, benzene, TX cut, Styrene andother C7-C8) is significantly higher when naphtha is processed accordingto the present invention (case 2) than can be achieved by conventionalprocessing means (case 1).

Because in the hydrocracking unit (case 2) the heavier paraffins are allreduced to lighter components such as C2-C4 paraffins, the production ofethylene increases from 35 to 50% by pre-treating the naphtha in thehydrocracking unit. Thus case 2 provides a significantly higher ethyleneyield than case 1.

Table 2 also shows that the production of heavier products (C9Resinfeed, cracked distillate and carbon black oil) is reduced bypre-treating the naphtha in the hydrocracking unit (case 2). This meansthat according to the present method the formation of heavy unwantedbyproducts can be reduced to a minimum.

Another example shows the influence of the operating temperature of ahydrocracking unit on the product slate. The catalyst mixture is aphysical mixture of 2 gram ZSM-5 and 2 gram Pt on alumina catalyst,wherein 0.4-0.8 mm SiC chips have been incorporated in the catalyst bedto ensure good approximation to plug flow and reduce axial/radialtemperature differences. Olefins6 naphtha was used as feedstock (seeTable 3).

TABLE 3 composition of feed stock Carbon number Naphthenes i-Paraffinsn-Paraffins Olefins Aromatics Total  3 <0.01 <0.01 <0.01 <0.01 <0.01<0.01  4 <0.01 0.22 0.51 0.92 <0.01 1.65  5 3.99 14.63 17.77 1.06 <0.0137.47  6 8.34 13.40 10.75 0.07 2.74 36.31  7 4.77 5.00 3.12 <0.01 0.9313.82  8 2.56 2.16 1.61 <0.01 0.83 7.16  9 1.36 1.27 0.73 <0.01 0.343.70 10 <0.01 0.28 0.53 <0.01 <0.01 0.81 11 <0.01 <0.01 0.06 <0.01 <0.010.06 Total 21.03 36.98 35.08 2.07 4.85 100.00

In this example the temperature was varied between 425° C. and 500° C.(WHSV=1, H:HC=3, 200 psig). The effluent composition is shown in Table4.

TABLE 4 effluent composition 500° C. 475° C. 450° C. 425° C. Methane12.43 7.45 4.42 2.28 LPG (C2-C4) 72.66 78.50 80.29 78.10 n-Paraffins(C5+) 0.08 0.24 0.58 1.22 i-Paraffins (C5+) 0.02 0.06 0.24 1.03 Olefins<0.01 <0.01 <0.01 <0.01 Naphthenes <0.01 0.04 0.43 0.43 Aromatics 14.8112.23 9.98 9.98 Total 100 100 100 100

The data from Table 4 show that higher temperatures result in highmethane (low value by-product) yields.

1. A method of processing a hydrocarbon stream that comprises middle distillate boiling range hydrocarbons, the method comprising: contacting the hydrocarbon stream with hydrogen under reaction conditions sufficient to produce a gaseous stream comprising at least liquefied petroleum gas; and separating the liquefied petroleum gas from the gaseous stream to produce a purified LPG stream; wherein the middle distillate boiling range hydrocarbons comprise kerosene, atmospheric gas oil, diesel, hydrocarbons from another hydrocracking unit having a boiling point in gasoil boiling point range, or combinations thereof; and wherein the reaction conditions comprise a temperature of 300-450° C., a pressure of 300-5000 kPa gauge and a Weight Hourly Space Velocity of 0.1-10 h⁻¹.
 2. The method of claim 1, wherein the contacting step further produces naphtha.
 3. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W, V, or combinations thereof in metallic or metal sulphide form supported on an acidic solid including alumina, silica, alumina-silica, zeolites, or combinations thereof.
 4. A method of processing a hydrocarbon stream that comprises middle distillate boiling range hydrocarbons, the method comprising: reacting hydrocarbons of a liquid hydrocarbon feedstock with hydrogen in the presence of a catalyst under reaction conditions sufficient to produce a gaseous stream comprising at least hydrogen, methane, liquefied petroleum gas (LPG), or combinations thereof; separating the gaseous stream into an LPG stream comprising predominantly liquefied petroleum gas, and an off-gas stream comprising hydrogen and/or methane; steam cracking the liquefied petroleum gas of the LPG stream in a steam cracking unit; and flowing the off-gas stream to the steam cracking unit to provide fuel for heating a furnace of the steam cracking unit.
 5. The method of claim 4, where in the steam cracking comprises: separating C2-C4 paraffins from said gaseous stream, further separating C2-C4 paraffins into individual streams, each stream predominantly comprising C2 paraffins, C3 paraffins and C4 paraffins, respectively; and feeding each individual stream to a specific furnace section of said steam cracker unit.
 6. The method of claim 4, wherein the reaction conditions comprise a temperature of 300-450° C., a pressure of 300-5000 kPa gauge and a Weight Hourly Space Velocity of 0.1-10 h⁻¹.
 7. The method of claim 4, where in the contacting step is conducted in the presence of a catalyst comprising Pd, Rh, Ru, Ir, Os, Cu, Co, Ni, Pt, Fe, Zn, Ga, In, Mo, W, V, or combinations thereof in metallic or metal sulphide form supported on an acidic solid including alumina, silica, alumina-silica, zeolites, or combinations thereof.
 8. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Pd.
 9. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Rh.
 10. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Ru.
 11. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Ir.
 12. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Os.
 13. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Cu.
 14. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Co.
 15. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Ni.
 16. The method of claim 1, where in the contacting step is conducted in the presence of a catalyst comprising the Pt.
 17. The method of claim 7, where in the contacting step is conducted in the presence of a catalyst comprising the Cu.
 18. The method of claim 7, where in the contacting step is conducted in the presence of a catalyst comprising the Co.
 19. The method of claim 7, where in the contacting step is conducted in the presence of a catalyst comprising the Ni.
 20. The method of claim 7, where in the contacting step is conducted in the presence of a catalyst comprising the Pt. 